Hydrocarbon conversion process



HYDROCARBON CONVERSION PROCESS Filed April 8, 1954 2 Sheets-Sheet 1 LeRoy v. Roms JR. Enum mcnuncu "VENTORS MMX/571417 monnfv July 21, 1959 L. v. RoBBlNs, JR., ETAL .2,895,900

HYDRocARBoN CONVERSION PRocEss Filed April 8,l 1954 2 Sheets-Sheet 2 www ATTORNEY United States Patent HYDROCARBON CONVERSION PROCESS Leroy V. Robbins, Jr., and Edwin J. Newchurch, Baton lRouge, La., assignors to Esso Research and Engineering Company, a corporation of Delaware Application April 8, 1954, Serial No. 421,852

2 Claims. (Cl. 208-72) This invention relates to the conversion of hydrocarbons to lower boiling components and more particularly relates to a process for conversion of hydrocarbons wherein the conversion to coke is minimized. Still more particularly the invention is concerned with a process which comprises a combination of catalytic and thermal cracking steps whereby a high ratio of gasoline to coke is effected.

Catalytic cracking of hydrocarbons is well known in the petroleum industry. This particular process has made it possible to produce gasoline from petroleum fractions, such as gas oils, which have a boiling range substantially above the gasoline boiling range. Because of this process it has been possible to substantially increase the proportion of gasoline obtained from crude oil. In addition the octane number of the gasoline produced by catalytic cracking is substantially higher than the octane number of the gasoline fraction of crude oil.

In the conventional catalytic cracking process, a hydrocarbon gas oil is contacted with a bed of catalytic material in a reactor to thereby convert the gas oil to lower boiling components. For this purpose a number of types of catalytic beds have been employed such as iixed beds,

moving beds and fluidized beds, all of which are well known in the petroleum industry. During the contacting of the gas oil with the catalyst, in any of the aforementioned types of catalyst beds, a carbonaceous or coke-like deposit is laid down on the catalyst. This carbonaceous deposit reduces the effectiveness of the catalyst to convert gas oil to lower boiling components and it isAtherefore necessary to remove this coke-likedeposit yfrom the catalyst, or in other words, to regenerate the catalyst so that the catalyst may be used for further catalytic cracking. This is accomplished by separating the coke-fouled or spent catalyst from the hydrocarbons and burning of the coke-like deposit with an oxygen-containing gas.v

In the case of catalytic cracking processes employing the catalyst in the form of a fixed bed, the introduction of gas oil to the catalyst bed `is discontinued when the catalyst becomes fouled with coke and the bed is then regenerated by introducing an oxygen-containing gas to the bed in order to burn olf the carbonaceous. or coke-like'deposit on the spent catalyst. After regeneration, the catalyst bed can be utilized for further catalytic cracking. In the case of catalytic cracking processes yemploying catalyst in the form of a moving bed or a fluidized bed, the spent catalyst is withdrawn continuously from the reactor without interrupting the cracking process. An oxygen-containing gas is then combined with the spent catalyst and the mixture is introduced into a vessel called a regenerator in which the carbonaceous deposits on the spent catalyst are burned olf. After regeneration the icatalyst is returned to the reactor and again employed to catalytically crack more gas oil.

By changing reaction conditions, such as temperature, space velocity, residence time, etc., the conversion of gas oil to gasoline may be increased in a catalyst cracking process. However, in general, asthe conversion of gas 2,895,900 Patented July 21, 1959 oil to gasoline is increased, the amount of coke formed and laid down on the catalyst also increases. Although it is desirable to maximize the amount of gas oil converted to gasoline, the extent of this conversion is thus limited by the amount of carbonaceous or coke-like deposits laid down on the catalyst and the facilities provided for burning o these carbonaceous deposits. Therefore, there is normally an economic balance between the conversion of gas oil to gasoline and the investment required for regeneration facilities. As regeneration facilities represent` a substantial portion of the investment for Ya catalytic cracking system, the conversion of gas oil to gasoline in a given catalytic cracking system is therefore normally limited to the capacity of the regeneration facilities which were originally economically justified. Certain gasoils such as those which contain a large proportion of aromatics or are high in nitrogen will produce a substantially higher ratio of coke to gasoline than will parainic or naphthenic gas oils. Because of the high coke forming tendencies of these gas oils, it is normally necessary to reduce the conversion of gas oil to gasoline since the regeneration facilities limit the catalytic cracking system. Heretofore there was no way in which a relatively high conversion to gasoline could be obtained from such gas oils without building abnormally high capacity regeneration facilities.

An object of the present invention is to effect a high conversion of gas oil to gasoline while at the same time to minimize the amount of coke formed on the catalyst'.

Another object of the present invention is to reduce the investment in regeneration facilities required for catalytic cracking systems.

Still further objects of the present invention will be apparent from a reading of this specification.

Briefly, the present invention comprises contacting a hydrocarbon gas oil, substantially all of which boils above about 430 F., with a cracking catalyst at a temperature of about S50-1000" F. and preferably at a temperature of about`900-950 F. so that between about C50-60%V by volume of the hydrocarbon gas oil is converted to gasoline, other components boiling below about 430 F. and coke and so that not more than about 5% by weight of the hydrocarbon gas oil is converted to coke. The resultant vaporous lower boiling hydrocarbon products are then separated from the resultant spent catalyst so that the catalyst concentration in the separated vaporous hydrocarbon products does not exceed more than about 0.1 lb. of catalyst/cubic foot. The substantially catalyst free vapors are then heated to a temperature of about 1050- 1l50 F. and preferably to a temperature of about 1075- 1125 F. and are maintained or soaked in that temperature range for about 0.5-l0 seconds so that about an additional 520% by volume of the vaporous hydrocarbons are converted to gasoline and other componentsboiling below about 430 F. During this soaking step of the present invention, substantially no coke Y is formed. Thereafter the hydrocarbon vapors are recovered and fractionated into a number of desired fractions including a gasoline fraction. Thus the initial contacting of the hydrocarbon gas oil with the catalyst results in the major portion of the conversion of the hydrocarbon gas oil to high-quality gasoline and the subsequent soaking step or predominantly thermal cracking step provides additional conversion with essentially no further coke formation. The present invention is applicable to hydrocarbon gas oils generally but is especially useful for high carbonproducing gas oils.

The process of the present invention will be readily understood by reference to the drawings in which:

Fig. l is a diagrammatic view of one form of apparatus adapted to canry out the process of the present invention.; and l Y Fig. 2 is a diagrammatic view of another form of apparatus adapted to carry out the process of the present invention.

It is to be understood that the present vinvention isY not limited to the particular apparatus shown inthe drawings as a reading of this specification will suggest to those skilled in the art other forms of apparatus for carrying out the process of the present invention.

Referring now to Fig. 1, reference character 10 designates a vessel or reactor adapted to carry out a catalytic conversion of a hydrocarbon gas oil with a finely divided uidized catalyst. In this specific embodiment of the present invention, hot freshly regenerated catalyst is introduced into the bottom of reactor 10 through inlet conduit 11 and its rate of introduction thereto is controlled by means of valve 12 in conduit 11. The catalyst employed may be a natural catalytic substance, such as acid-treated clay, or asynthetic catalytic substance such as silica-alumina, silica-magnesia, etc. In general the fine- 1y divided catalyst particles have a size of less than about 250 microns in average diameter and normally the particles will have a size range such that substantially all of them are between about 20-100 microns in average diameter. The catalyst flowing through conduit 11 will normally be obtained from a regenerator (not shown) wherein carbonaceous deposits are burned olf the spent catalyst produced in the present invention as will hereinafter be described in detail. Thus in the specific embodiment of the present invention as shown in Fig. 1 the finely divided catalyst will be continuously passed between reactor 10 and the regenerator which may be any conventional regenerator adapted to regenerate a finely divided catalyst.

Prior to the introduction of the finely divided catalyst into the bottom of reactor 10, it is combined with a hydrocarbon gas oil introduced into conduit 11 through conduit 13 at a rate regulated by means of valve 14 in conduit 13. The hydrocarbon gas oil will generally have a boiling range between about 400-l050 F. Normally the hydrocarbon gas oil introduced through conduit 13 will have been preheated to a temperature which may be as high as about 700 F. -by indirect heat exchange with high temperature product streams or by a preheat furnace (not shown). The amount of preheat required will depend upon a number of factors such as the temperature of the regenerated catalyst in conduit 11 which may be about l000`l150 F., the catalyst-to-oil feed ratio to vessel 10 which may be from about 5-15 by weight, the reaction temperature desired in vessel 10, etc.

The catalyst-oil mixture is then introduced into the bottom of reactor 10 at a rate such that the superficial velocity of the oil vapors rising upwardly therein will be about 0.5- ft./sec. and preferably about 1 3 ft./sec. At these vapor velocities the finely divided catalyst will form a dense iluidized bed 15 in the lower portion of vessel with the .bed having a density of about 30-50 lbs.v/cu. ft. The reaction temperature in vessel 10 will be maintained at about 850-1000 F. and preferably at about 900-950 F. and the pressure in reactor 10 will normally be maintained at about 0-50 lbs/sq. in. gauge (p.s.i.g). As previously stated, the finely divided catalyst in reactor 10 will fbe generally maintained as dense bed in the lower part of reactor 10 and this dense bed will have a relatively well-defined upper level L which is maintained thereat by continuously .withdrawing catalyst from dense bed 15 by means of outlet conduit 20 which extends through vessel 10 up into dense bed 15 at a position slightly below upper level L. The rate of withdrawal of catalyst from dense bed v15 through conduit is regulated by means of valve 21 in conduit 20. The catalyst hold-up in reactor 10 should be sufiicient to provide a space velocity of about 1-5 lbs. of hydrocarbon oil/hour/ lb. of catalyst hold-up in reactor 10. The catalyst removed from dense bed 15 through conduit 20 is passed to the previously mentioned regenerator (not shown) associated with reactor 10 wherein carbonaceous deposits are burned olf the spent catalyst with air, after which the resultant regenerated catalyst is passed back to reactor 10 through conduit 11.

Thus in the specific embodiment of the present invention shown in Fig. 1, the hydrocarbon vapors are passed upwardly in reactor 10 through dense bed 15 wherein they are converted to lower boiling products and coke which is deposited on the finely divided catalyst to produce spent catalyst. The reaction conditions in reactor 10 are adjusted such that about 30-50% of the hydrocarbon gas oil introduced into the bottom of reactor 10 is converted to products boiling below about 430 F. and coke, and so that no more than about 5% by weight of the hydrocarbon gas oil is converted to coke. After flowing through dense bed 15, the converted hydrocarbon vapors enter dilute or disperse phase 22 located in reactor 10 above upper level L of dense bed 15. Disperse phase 22 contains a small amount of catalyst which is entrained with the vapor rising from dense bed 15. The rising hydrocarbon vapors in disperse phase 22 pass upwardly out of reactor 10 through outlet conduit 23. These vapors are essentially free of entrained catalyst and contain no more than about 0.1 lb. of catalyst/cu. ft. of hydrocarbon vapors, and normally contain only about 0.002-0.05 lb. of catalyst/ cu. ft. of vapors.

The hydrocarbon vapors passnig upwardly through conduit 23 from reactor 10 are introduced into heat exchanger 30 wherein they flow through the interior of tubular heating coil 31. In heat exchanger 30, the hydrocarbon product vapors are heated to a temperatures of about 1050-l F. and preferably to a temperature of about 1075-1125 F. It is to be clearly understood that all of the vaporous hydrocarbon products formed in reactor 10 are heated to this particular temperature, name ly 50-300 F. above the temperature existing in reactor 10. Heat exchanger 30 may, for example, comprise a furnace into which is introduced a `gas or oil fuel through conduit 32 containing valve 33 and an oxygen-containing gas such as air through conduit 34 containing valve 35. The rates of introduction of oil and air through conduits 32 and 34, respectively, are controlled by means of valves 33 and 35, respectively, to produce a combustible mixture which is burned in the interior of heat exchanger or furnace 30. The gaseous products of combustion Afrom heat exchanger 30 are removed therefrom through conduit 37.

Instead of utilizing heat exchanger 30 as a furnace, it may be practical to introduce the hot ue gase from the regenerator associated with reactor 10 into the interior of heat exchanger 30 through conduits 32 and 34. Thus if the temperature of the regenerator ue gas is about 1100 F. or above, it may be employed in heat exchanger 30 to heat the product vapors passing through coil 31. In this case, conduit 37 may be employed to continuously remove the cooled regenerator flue gas from exchanger 30. '111e gases exiting from heat exchanger 30 through conduit 37 may be vented directly to the atmosphere or may be passed through other heat exchangers to recover a major portion of their sensible heat. For example, the hot gases passing from heat exchanger 30 through conduit 37 may be employed to preheat the fresh hydrocarbon gas oil which is introduced to the system of the present invention through conduit 13.

The length and diameter of coil 31 are selected so that the vaporous hydrocarbon products which are introduced thereto from conduit 23 will be heated to a temperature of about 1050-1150" F., and preferably to a temperature of about 1075-1125 F., because it will be necessary to maintain the heated vapors in this temperature range for aboutvO.5-10 seconds prior to their being fractionated into desired fractions as will hereinafter be described in detail. The particular length and diameter of coil 31 required will be dependent upon the hydrocarbon vapor asssgooo flow rate, the temperature of the heating medium in heat exchanger 30, the heat exchange coeicients involved, etc. The temperature to which the hydrocarbon vapors are heated in coil 31 and the residence time of the vapors in the temperature range of l050-1150 F. prior to fractionation should be selected such that an additional 5-20% of the hydrocarbon vapors are converted to products boiling below about 430 F. The required residence time of the vapors in the specied temperature range will depend of course upon the temperature involved; thus thevhigher the temperature, the shorter the residence time, and vice versa. Thus flexibility of conversi-on may be readily attained by regulating the extent of heating in coil 31. The additional amount of converted hydrocarbons will comprise mainly additional gasoline and dry gas(C3 and lighter gases). In addition to these products a very small amount of the vaporous hydrocarbons introduced into coil 31 will be converted to coke. Normally this amount of coke will not exceed more than about 0.5% by weight of the original hydrocarbon gas oil fed to reactor through conduit 13.

The heated hydrocarbon vapors are removed from coil 31 by means of conduit 40 which ycommunicates at its other end with the interior of cyclone separator 41. The heated hydrocarbon vapors which are being thermally cracked are introduced tangentially into housing 42 of cyclone separator 41 through inlet 43. In housing 42 the vapors are swirled around therein such that the small amount of entrained catalyst is thrown against the inner wall of housing 42 due to centrifugal force to separate it from the bulk of the hydrocarbon vapors therein. The catalyst falls downward along the inner wall of housing 42 and flows downwardly from housing 42 through dipleg 45 which communicates at its lower end with dense bed in reactor 10. Thus the entrained finely divided catalyst is separated from the hydrocarbon vapors in cyclone separator 41 and is returned to the system by means of dipleg 45 which extends at its lower end beneath upper level L of dense bed 15. The separated hydrocarbon vapors are removed from cyclone separator 41 through outlet conduit 44 which is arranged centrally in housing 42 and which communicates with the upper portion thereof. It will be understood that the positions of cyclone separator 41 and heat exchanger 30 may be interchauged so that the hydrocarbon vapors from reactor 10 would be passed through cyclone separator 41 first and then through heat exchanger 30 after which the resultant heated vapors would be passed into conduit 44. In this case cyclone separator 41 could be arranged within the upper portion of reactor 10, if desired.

The hot hydrocarbon vapors passing through conduit 44 are introduced into the lower portion of fractionator 50 wherein the hydrocarbon products are partially cooled by retluxing and are separated into at least two fractions. It will be understood that the residence time of 0.5-10 seconds required for soaking the heated hydrocarbon vapors in the temperature range of 1050-1150 F. extends from coil 31 to fractionator 50. Cyclone separator 41 and conduits 40 and 44 are preferably insulated in order to maintain the hydrocarbons at this soaking temperature. In this specific embodiment of the present invention, gasoline and other components boiling below about 430 F. are removed overhead from fractionator 50 through conduit 51 and components boiling above about 430 F. are removed as a bottoms fraction through conduit 52. It will be understood however that the hydrocarbon products may be fractionated into more product streams if desired by providing fractionator 50 with sidestream drawolfs (not shown). Fractionator 50 may be any conventional fractionating means for separating hydrocarbons and thus may be provided with bubble-cap trays, jet trays, etc. to increase the fractionating eiciency. If desired, a cooling means may be installed in the apparatus shown in Fig. l in order to'partially cool the vapors from heat exchanger 30 before the vapors are introduced into cyclone 'separator 41. Thus for example the vapors could be cooled to about 850- 900 F. prior to their introduction into cyclone separator 41. In this event, however, the vapors should be maintained in exchanger'30 at a temperature of about 1050-. 1150 F. for 0.5-10 seconds so that the additional 5-20% conversion will be obtained. This may be readily accomplished by increasing the capacity of coil 31 so that the residence time of the vapors in heater 30 is increased suciently to accomplish the desired conversion; Thus it will readily be seen that there are many forms of apparatus which may be employed to carry Yout the method ofthe present invention. l

Referring now to Fig. 2, reference character designates an elongated vessel which is adapted to carry outa hydrocarbon conversion' reaction in the presence of-`a Afinely divided catalyst and which will be termed transfer line reactor 100. A nely divided uidized catalyst is introduced into the bottom of transfer line reactor 100 through conduit 101 containing valve 102. The finely divided catalyst n this specific embodiment of the present invention has a size of less than about 250 microns in average diameter and generally all of the nely divided catalyst will .have a size range of about 20-100 microns. The nely divided catalyst passing through conduit 101 is normally a hot vfreshly regenerated catalyst obtained from a regenerator (not shown) wherein carbonaceous deposits have been burned otf of spent catalyst which is produced in the process of the present invention as will be hereinafter described in detail.

A hydrocarbon gas oil having a boiling range between 40o-1050" F. is introduced into conduit 101 through conduit 103 containing valve 104 for admiXture therein with the finely divided catalyst. This hydrocarbon gas oil may be preheated to a temperature as high as about 700 F. by means of heat exchangers (not shown) or by a preheat furnace (not shown). The relative amounts of catalyst and hydrocarbon gas oil introduced into the bottom of reactor 100 are'regulated by means of valves 102 and 104 in conduits 101 and 103, respectively. The rate of introduction of the ygas oil-catalyst mixture in transfer line reactor 100 is adjusted so that the Yupward superficial velocity of the rising hydrocarbon vapors therein will be about 6-15 ft./sec. and so that the catalyst concentration in transfer line reactor 100 will be about 10-20 lbs/cu. ft. Underl .these conditions, the hydrocarbon vapors and catalyst will ow upwardly in transfer line reactor 100 substantially concurrently such that no dense bed of finely divided catalyst having a well-defined upper level 'as in thel specific embodiment `shown in Fig. l will be formed in the bottom of reactor 100. In flowing upwardly through transfer line reactor 100, the vaporous hydrocarbon gas oil at a temperature of about 850-1000 F. is converted to about 30-50% by volume of products boiling 'below about 430 F. and coke which is deposited on the finely divided catalyst. The amount of coke formed -should not exceed more than about 5% by weight of the original hydrocarbon gas oil.

The resultant suspension of lower boiling vaporous hydrocarbon products and spent catalyst `are withdrawn from the top of transfer line reactor 100 through con- -duit which communicates at its other end with cy clone separator 111. The vaporous hydrocarbon products and catalyst are introduced tangentially into housing 112 of cyclone separator 111 through inlet 113. The catalyst-oil suspension is swirled around in housing'112 to effect a rapid separation of the nely divided catalyst from the hydrocarbon vapors which are removed from the upper part of housing 112 through centrally located outlet pipe 114. The separated catalyst particles ow downwardly out of housing 112 through conduit 115 wherefrom the separated finely divided catalyst is passed to a regenerator (not shown) so as to burn oi the carbonaceous deposits from the spent catalyst. y &1bsequently thisregenerated catalyst is reintroduced into transfer line' reactor 100 from the regenerator through conduit 101 for contact with additional hydrocarbon gas oil.

The vaporous hydrocarbon products which are substantially freed of the finely divided catalyst in cyclone separator 111 are passed therefrom through conduit 114 into heat exchanger 120 wherein they ow through the interior of tubular heating and soaking coil 121. Heat exchanger 120 may, for example, comprise a furnace wherein a liquid or gaseous fuel is introduced to the interior thereof through line 122 containing valve 123 and wherein an oxygen-containing gas such as -air is introduced to the interior thereof through line 124 containing valve 125, Valves 123'and 125 in lines 122 and 124, respectively, are regulated so as to produce in the interior. of h eat exchanger 120 a: combustible mixture whichis burned to etfect'the heating desired in this invention. In this case the products of combustion are removed from the interior of heat exchanger 120 through outlet pipe 126. AThe gaseous products of combustion exiting from heat exchanger 120 through outlet line 126 may be vented directly to the atmosphere or may be first passed through'other heat exchange equipment (not shown) to recover the major portion of their sensible heat. Alternatively heat exchanger 120 may employ hot flue gases having a temperature of about 1050- 1l50 F. which are produced in the regenerator which is associated with transfer line reactor 100. In this case, the hot liue gases from the regenerator would be introduced through line 123 and/or line 124 and would be removed from the interior of heat exchanger 120 through outlet line 126.

The diameter and length of coil 121 are selected so that the vaporous hydrocarbon products passing through coil 121 are heated therein to a temperature of about 1050-1150" F. and preferably to a temperature of about 1075-1125 F. The dimensions of coil 121 which are required will depend upon the heat transfer coelicients involved, the temperature of the heating medium, etc. Also it is necessary thatA the resultant heated vapors be maintained in this temperature range for about 0.5-10 seconds prior to being fractionated into desired fractions as will hereinafter be described in detail so that an Aadditional 5-20% of the original hydrocarbon gas oil is' converted to products boiling below 430 F. and a very small amount of coke which will not exceed more than about 0.5% by weight of the original hydrocarbon gas oil. It is to be clearly understood that all of the vaporous hydrocarbon products formed in reactor 100 are heated to this particular temperature, namely 50- 300 F. above the temperature existing in reactor 100. The particular temperature to which the hydrocarbon vapors are heated in heat exchanger 120 and their residence time in the temperature range of about 1050- 1150 F. prior to fractionation will of course be dependent upon the thermal cracking properties of the particular hydrocarbon gas oil introduced into the system of the present invention.

The heated vaporous hydrocarbons which are being thermally cracked are removed from coil 121 of heat exchanger 120 by means of conduit 130 whichy communicates at its other end with fractionator 131 wherein the hydrocarbon products are partially cooled by refluxing and are separated into at least two product streams. In this particular embodiment of the present invention, the converted hydrocarbons are separated into a fraction boiling below about 430 F. which is taken off overhead through conduit 132 and a bottoms fraction which is withdrawn through conduit 133. However, it will be understood that additional product streams may be taken by providing fractionator 131 with sidestream drawoffs (not shown). Fractionator 131, which is of conventional design, may be provided with bubble cap trays, jet trays, etc. in order to improve its fractionating eiiiciency. If desired, the thermally cracked vapors from heat exchanger may be cooled or quenched below an active cracking temperature before they are introduced into fractionator 131 by passing the vapors through heat exchangers or quenching towers (not shown). However, it is essential that the vapors prior to this cooling or quenching step be maintained in the temperature rangeof about 1050-1l50 F. for 0.5-10 seconds so that the desired additional conversion is obtained,

A reading of this specification Will suggest to those skilled in the art other forms of apparatus for carrying out the process of the present invention and it is to be clearly understood that the present invention is not limited to the particular apparatus shown in the drawings. It is to be further understood that the present invention is not'limited solely to fluidized solids systems such as the dense fluid bed type shown in Fig. l and the transfer line reactor type shown in Fig. 2. Thus, more specifically the process of the present invention is also ideally suited for utilization in a system in which the catalyst is maintained either as a fixed bed or as a moving bed. In either of these two latter types of systems, the present process could be accomplished by passing the hydrocarbon gas oil vapors through the catalyst bed under cracking conditions of temperature (850-1000 F.) and pressure (0-50 p.s.i.g.) such that about 30-50% of the hydrocarbon gas oil would be converted to products boiling below 430 F. and a small amount of coke not exceeding more than about 5% by weight of the original hydrocarbon gas oil.

The resultant vaporous reaction products from the catalytic cracking step would then be heated to about 1050-1150 F. and maintained in this temperature range for about 0.5-10 seconds to convert an additional 5-20% of the original hydrocarbon gas oil to products boiling below 430 F. anda small amount of coke not exceeding 0.5% by weight of the original hydrocarbon gas oil, after which the resultant further converted hydrocarbon vaporous products would be separated into desired product fractions. When the method of the present invention is employed with a fixed catalyst bed or a moving catalyst bed, the present invention would be somewhat simplified in that there would be essentially no problem involved in the entrainment of catalyst in the vaporous hydrocarbon products as is encountered in the fluidized systems. However, the fluidized systems generally exhibit certain advantages over the fixed bed and moving bed types of system such as better utilization of heat from the regenerator, easier control of process variables, etc.

The following experiment was carried out and is set forth hereinafter in order to illustrate the advantages to be derived from the method of the present invention as compared to the conventional catalytic cracking process. The experimental data was obtained from a 68 gram batch uid catalytic cracking unit in which an Elk Basin heavy gas oil, which is a typical catalytic cracking feed stock, was catalytically cracked with a commercial silicamagnesia catalyst having a size range of about 44 to 147 'microns The Elk Basin heavy gas oil had a boiling range of about 400 to l050 F. and a gravity of about 23 A.P.I. In general, the finely divided silica-magnesia catalyst was maintained during this experiment in the bottom ofthe reactor of the catalytic cracking unit in the form of a dense fluidized bed. Above this dense uidized bed there was a dilute phase or vapor space in which the concentration of catalyst was about 0.01 lb./cu. ft. This dilute phase was provided with an electrical heating means so that the temperature in the dilute phase could be maintained at a higher temperature than that existing in the dense bed. The vaporous hydrocarbon products were passed from the cracking unit through a water jacketed condenser and were collected in a receiver placed in a Varsol-Dry Ice bath. The condensed liquid hydrocarbon products were then cut in a small batch' still to determine product yields.

In-the first partof this experiment, the dense bed and the dilute phase were maintained at the same temperature, namely, 925 F., and the Elk Basin gas oil was catalytically cracked under the conditions set forth in the table below. After the gasoline and carbon yields and conversion had been determined for these operating conditions, the operating conditions were changed so that in the second part of the experiment the dense bed was maintained at about 925 F. and the dilute phase or vapor space above the dense bed was maintained at l100 F. by employing the previously mentioned heating means. Again the carbon and gasoline yields and conversion were determined for this second set of operating conditions. The following results were obtained for (1) the conventional catalytic cracking method and (2) the catalyticthermal cracking method of the present invention:

Pressure, P.s.i.g. (Both Zones) Conversion, Vol. Percent of Feed Carbon, Wt. Percent of Feed- Gasoline, Wt. Percent of Feed- 31 39.

Gasoline to Carbon Ratio 7. 5 8.7.

Carbon Yield at Comparable Conversion 5.1 2.6.

Levels (45).

It will be noted from the experimental data that the percent conversion (percent of gas oil feed converted to products boiling below 430 F. and coke) obtained with the method of the present invention was substantially higher than that obtained with the conventional catalytic cracking method under comparable operating conditions. More speciiically, the conversion obtained in accordance with the present invention was 14% higher than that of the conventional catalytic cracking method for comparable operating conditions. It will be further noted that this substantial increase in conversion was realized with only a slight increase in the amount of carbon produced. This increase in conversion was in part reiiected by an 8% increase in the amount of gasoline produced. 'I'hus it will be seen that the gasoline to carbon ratio for the present invention was substantially higher than that of the conventional catalytic cracking process, namely 8.7 as compared to 7.5.

The advantage of the method of the present invention is even more emphatically brought out by the last set of figures presented in the above table wherein the carbon yields of the two methods were compared at the same conversion level, namely, 45%. Thus it will be noted that'at the same conversion level the method of the present invention produces only about one-half as much carbon as does the conventional catalytic cracking method. This substantial reduction in carbon is significant as this means that when employing the process of the present invention it is possible to employ a catalyst regenerator of only about half the capacity of that required in a conventional catalytic cracking process. Thus it is apparent that the method of the present invention will make possible substantial savings in the investment required for the regeneration portion of a catalytic cracking system. Itis to be understood, however, that the improved results evidenced by the present invention may be realized either as an increased gasoline yield or as a decreased carbon yield, or as a combination of the two.

What is claimed is:

1. A method for converting a hydrocarbon gas oil having a boiling range between about 400-1000" F. to lower boiling components including gasoline which comprises combining said hydrocarbon gas oil with a hot finely divided catalyst, introducing the resultant catalyst and vaporous hydrocarbon gas oil mixture at a temperature of about 8501000 F. into the bottom of a reaction Zone wherein said catalyst is maintained in the lower portion thereof as a dense fluidized bed by passing said vaporous hydrocarbon gas oil upwardly therethrough at a superficial velocity of about l to 3 feet/second and wherein about 3050% of said vaporous hydrocarbon gas oil is converted to coke and gasoline boiling below 430 F., but not more than about 5% by weight of said vaporous hydrocarbon gas oil is converted to coke and wherein the resultant converted hydrocarbon vapors are passed from said dense bed upwardly through a disperse phase in said reaction zone located above said dense bed to thereby disengage substantially all of said finely divided catalyst from said converted hydrocarbon vapors, introducing only said total converted hydrocarbon vapors containing less than about 0.1 lb. of entrained catalyst/ cubic foot into a soaking zone wherein said partially converted vaporous hydrocarbons are heated solely by indirect heat exchange to about l050-1150o F., maintaining said converted hydrocarbon vapors in said soaking zone at a temperature of about 1050-1150 F. for about 0.5-10 seconds until an additional 5-20% of said hydrocarbon vapors is converted to gasoline boiling below 430 F. and a small amount of coke, and passing the resultant further converted hydrocarbon vapors from said soaking Zone into a recovery zone wherein said further `converted hydrocarbon vapors are fractionated into at least two product fractions.

2. A method for converting a hydrocarbon gas oil having a boiling range between about 400-1000 F. to lower boiling components including gasoline which comprises combining said hydrocarbon gas oil with a hot nely divided catalyst, passing the resultant catalyst and vaporous hydrocarbon gas oil mixture at a temperature of about 850-1000n IF. upwardly through a transfer line reaction zone at a superficial velocity of about 6-15 feet/second to thereby convert about 30-5 0% of said vaporous hydrocarbon gas oil to coke and gasoline boiling below 430 F., but to convert no more than about 5% by weight of said vaporous hydrocarbon gas oil to coke, passing the resultant mixture of converted hydrocarbon vapors and spent catalyst from said reaction zone into a separation zone wherein substantially all of said spent catalyst is separated from said converted hydrocarbon vapors, introducing only the resultant separated hydrocarbon vapors from said separation zone into a soaking zone wherein said vapors are heated solely by indirect heat exchange to about l050-1150 F., maintaining said vapors in said soaking Zone at a temperature of about 1050-1150 F. for labout 0.5-10 seconds until an additional 5-20% of said hydrocarbon vapors is converted to gasoline boiling below 430 F. and a small amount of coke, and passing the resultant further converted hydrocarbon vapors from said soaking zone into a recovery zone wherein said further converted hydrocarbon vapors are fractioned -into at least two product fractions.

References Cited in the file of this patent UNITED STATES PATENTS Sachanen: Conversion of Petroleum (1940), page 111. 

1. A METHOD FOR CONVERTING A HYDROCARBON GAS OIL HAVING A BOILING RANGE BETWEEN ABOUT 400-1000* F. TO LOWER BOILING COMPONENTS INCLUDING GASOLINE WHICH COMPRISES COMBINING SAID HYDROCARBON GAS OPIL WITH A HOT FINELY DIVIDED CATALYST, INTRODUCING THE RESULTANT CATALYST AND VAPOROUS HYDROCARBON GAS OIL MIXTURE AT A TEMPERATURE OF ABOUT 850-1000* F.INTO THE BOTTOM OF A REACTION ZONE WHEREIN SAID CATALYST IS MAINTAINED IN THE LOWER PORTION THEREOF AS A DENSE FLUIDIZED BED BY PASSING SAID VAPOROUS HYDROCARBON GAS OIL UPWARDLY THERETHROUGH AT A SUPERFICIAL VELOCITY OF ABOUT 1 TO 3 FEET/SECOND AND WHEREIN ABOUT 30-50% OF SAID VAPOROUS HYDROCARBON GAS OIL IS CONVERTED TO COKE AND GASOLINE BOILING BELOW 430* F., BUT NOT MORE THAN ABOUT 5% BY WEIGHT OF SAID VAPOROUS HYDROCARBON GAS OIL IS CONVERTED TO COKE AND WHEREIN THE RESULTANT CONVERTED HYDROCARBON VAPORS ARE PASSED FROM SAID DENSE BED UPWARDLY THROUGH A DISPERSE PHASE IN SAID REACTION ZONE LOCATED ABOVE SAID DENSE BED TO THEREBY DISENGAGE SUBSTANTIALLY ALL OF SAID FINELY DIVIDED CATALYST FROM SAID CONVERTED HYDROCARBON VAPORS, INTRODUCING ONLY SAID TOTAL CONVERTED HYDROCARBON VAPORS CONTAINING LESS THAN ABOUT 0.1 LB OF ENTRAINED CATALYST/ CUBIC FOOT INTO A SOAKING ZONE WHEREIN SAID PARTIALLY CONVERTED VAPOROUS HYDROCARBONS ARE HEATED SOLELY BY INDIRECT HEAT EXCHANGE TO ABOUT 1050-1150* F. FOR ABOUT 0.5-10 SECONDS UNTIL AN ADDITIONAL 5-20% OF SAID HYDROCARBON VAPORS IS CONVERTED TO GASOLINE BOILING BELOW 430* F. AND A SMALL AMOUNT OF COKE, AND PASSING THE RESULTANT FURTHER CONVERTED HYDROCARBON VAPORS FROM SAID SOAKING ZONE INTO A RECOVERY ZONE WHEREIN SAID FURTHER CONVERTED HYDROCARBON VAPORS ARE FRACTIONATED INTO AT LEAST TWO PRODUCT FRACTIONS. 